Process for desulfurizing hydrocarbon fuels and fuel components

ABSTRACT

Processes are disclosed for removing sulfur, including cyclic and polycyclic organic sulfur components such as thiophenes and benzothiophenes, from a hydrocarbon feedstock including fuels and fuel components. The feedstock is contacted with a regenerable sorbent material capable of selectively adsorbing the sulfur compounds present in the hydrocarbon feedstock in the absence of a hydrodesulfurization catalyst. In one embodiment, the sorbent can be an active metal oxide sulfur sorbent in combination with a refractory inorganic oxide cracking catalyst support. In another embodiment, the sorbent can be a metal-substituted refractory inorganic oxide cracking catalyst wherein the metal is a metal which is capable in its oxide form, of adsorption of reduced sulfur compounds by conversion of the metal oxide to a metal sulfide. The processes are preferably carried out in a transport bed reactor.

FIELD OF THE INVENTION

The present invention relates to the desulfurization of hydrocarbons,particularly hydrocarbon fuels and hydrocarbon fuel components and theirprecursors. More particularly, the present invention relates to removalof sulfur, primarily organic sulfur, contaminants including organicsulfides, disulfides, mercaptans, thiophenes, benzothiophenes, anddibenzothiophenes, from hydrocarbon fuels such as gasoline, dieselfuels, aviation fuels, and from components and precursors of such fuelssuch as FCC naphtha, i.e., naphtha from a fluid catalytic cracker (FCC),FCC light cycle oil, coker distillate, and the like.

BACKGROUND OF THE INVENTION

Currently available gasoline contains sulfur contaminants at an averagecumulative level exceeding 300 parts per million by weight (ppmw) ofsulfur (i.e., calculated based on sulfur weight). On-road applicationdiesel fuel has a higher sulfur content ranging typically from 300 to2,000 ppmw. Combustion of gasoline and diesel fuels during use ininternal combustion engines, in turn, converts the sulfur contaminantsinto sulfur oxides. The sulfur oxides are environmentally undesirableand also have been found to have a long-term deactivation impact onautomotive catalytic converters that are used to remove nitrogen oxideand unburned hydrocarbon contaminants from automotive emissions.

In order to improve air quality, environmental protection agencies ofvarious industrialized countries have therefore announced or proposednew regulations requiring reduction in sulfur content of gasoline,diesel, and other motor fuels. In the United States, the EnvironmentalProtection Agency (EPA) is requiring that the sulfur content ofgasolines be reduced to a maximum of 30 ppmw by the year 2005 underrecently implemented Tier 2 regulations. Similarly, the EPA has enactedregulations to bring down the sulfur levels in diesel fuel used foron-road application to 15 ppmv or below by 2006. It is anticipated thatdue to public demand for a cleaner environment, the future will bringcalls for even stricter sulfur oxide emissions and fuel specifications;and, as a result, fuels containing nearly zero sulfur levels are beingdiscussed. Accordingly, the new regulations will require sulfurreduction of typically 90% or more by 2005, and perhaps complete sulfurremoval thereafter. At the same time, the sulfur content of commerciallyavailable crude oils produced in the United States and in neighboringAmerican countries has been generally increasing; thus the newregulations will require more drastic sulfur reduction in the future.Further reductions meeting nearly zero sulfur levels required byexpected future regulations will exacerbate this problem further.

Various technologies are currently available or have been proposed whichare believed to be capable of reducing sulfur contaminants in gasolineto 30 ppmw or less. According to a recent study conducted by EPA, theseavailable and proposed technologies include hydrotreating andadsorption-based processes (see Regulatory Impact Analysis—Control ofAir Pollution From New Motor Vehicles: Tier 2 Motor Vehicle EmissionsStandards and Gasoline Sulfur Control Requirements, EPA 420-R-99-023,United States Environmental Protection Agency, December 1999, ChapterIV, pp. IV-42--IV-65).

As detailed in the EPA study, the sulfur content of current gasolines isattributable primarily to fluidized catalytic crackers (FCC), or tocoker units, which convert heavy boiling stocks to gasoline componentsor precursors, i.e., naphthas. It has been reported that more than 90%of the sulfur in gasoline comes from streams produced in the FCC unit.The sulfur content of FCC naphtha varies from 150 to 3,000 ppmwdepending upon the sulfur concentration of feed and the endpoint of thegasoline product. Accordingly, reduction of sulfur in motor gasoline canbe accomplished by FCC feed hydrotreating or by hydrotreating thenaphtha cut obtained from the FCC unit. The latter process is preferredbecause of substantially lower cost resulting from substantially lowervolumes of the feedstocks to be processed.

Nevertheless, hydrotreating of FCC naphtha is expensive, both in capitalinvestment, and in operating costs. In particular, hydrotreating of FCCnaphtha is typically carried out in a packed-bed or a fixed-bed reactorusing various well-known hydrodesulfurization (HDS) catalysts. Thesecatalysts typically contain a Group 8 (other than iron), 9, or 10transition metal such as cobalt and/or nickel combined with a Group 6transition metal, particularly molybdenum or tungsten, on a high surfacearea alumina support (“Group metal” as used herein is based on the newIUPAC format for the Periodic Table of the Elements, which numbers thegroups from 1 to 18 in Arabic numerals). Before their use, thesecatalysts are typically pre-sulfided under controlled reducingconditions to impart their HDS catalytic activity. Other HDS catalystsinclude platinum, palladium, or like metals supported on alumina. In thepresence of HDS catalysts, organic sulfur compounds present in FCCnaphtha react with hydrogen and are converted into hydrogen sulfide attemperature and pressures or 300 to 500° C., and 400 to 600 psig. Thehydrogen sulfide thus formed can be subsequently and readily removed ina downstream unit by sorbents or other processes such as a combinationof amine and Claus processes.

However, during the HDS hydrotreating process, octane number loss canoccur by saturation of high-octane containing olefins that are presentin FCC naphtha. Moreover, increased olefin saturation is accompanied byincreased hydrogen consumption and cost. In addition, there can be aloss in gasoline yield caused by mild cracking which breaks some of thenaphtha into smaller, lighter fractions, which are too light forblending into gasoline.

Three proven hydrotreating desulfurization technologies are identifiedin the EPA report cited previously. However, octane number loss remainsa serious problem with all three proven technologies particularly whenapplied for removal of 90 percent or more sulfur from the FCC naphtha tomeet EPA's Tier 2 requirements.

Newly proposed technologies identified in the EPA report include acatalytic distillation technology, called CDTech, which relies upon anHDS catalyst supported in a distillation column to provide reaction oforganic sulfur compounds with diene compounds present in FCC naphtha.The resultant thioether reaction product has a higher boiling point andcan be removed from the bottom of the distillation column. Similar toconventional hydrotreating processes, this process also uses an HDScatalyst. However, hydrogen consumption and olefin saturation areclaimed to be lower compared to conventional hydrotreating processes.The operating cost for sulfur removal using the CDTech process isreported to be 25% lower than conventional hydrotreating processes forthe same degree of sulfur removal.

Two emerging adsorption-based desulfurization processes are alsodiscussed in the EPA report. One process, named IRVAD, adsorbsheteroatom-containing hydrocarbon compounds, including sulfur, nitrogen,and oxygen compounds, present in FCC naphtha onto an alumina-basedadsorbent in liquid phase (see U.S. Pat. No. 5,730,860, issued Mar 24,1998 to Irvine). The adsorbent is fluidized in a tall column andcontinuously removed and regenerated using hydrogen in a second column.The regenerated catalyst is then recycled back into the reactor. Theregeneration of spent adsorbent produces a hydrocarbon stream containingabout 1 wt % sulfur, which can be treated using conventional processes.While the inventors have claimed an overall cost of sulfur removal aslow as 0.77 cents per gallon of gasoline compared to 5 to 8 cents forconventional hydrotreating processes, serious process and systemintegration issues still remain with this technology, which arehampering its commercial deployment.

The other emerging adsorption-based desulfurization technology named asthe SZorb process is being developed by the Phillips Petroleum Company.It is understood that this process uses an adsorbent/catalyst comprisingone or more metallic promoters, such as a combination of nickel andcobalt, in a zero valence state to selectively remove sulfur compoundsfrom FCC naphtha in the presence of hydrogen. As the adsorbent/catalystbecomes saturated with sulfur compounds, it is sent to a regenerationunit where it is treated with an oxygen-containing gas for removal ofthe sulfur as sulfur dioxide. The oxidized adsorbent/catalyst is furthertreated with hydrogen in a downstream reducing unit presumably to reducesome of the metal oxide/s present in the adsorbent/catalyst compositionto their reduced forms. The reduced adsorbent/catalyst is then fed tothe sulfur removal unit, along with hydrogen, for furtherdesulfurization of FCC naphtha. This process is carried out at atemperature between about 250 to about 350° C. (about 500 to about 700°F.) and a pressure of 100 to 300 psig. Phillips proposes to useconventional bubbling-bed fluidized-beds for adsorption and regenerationreactors, which will have inherent limitation on throughput of the FCCnaphtha feed that can be processed in this system. Phillips claims thatthis process can remove about 97% of the sulfur from FCC naphtha with a1 to 1.5 point loss in octane number and with an operating cost of 1.5to 2 cents per gallon of gasoline. However, the need for a two-stepregeneration process, consumption of hydrogen and associated octanenumber loss, and the use of low throughput bubbling-bed systems are someof the major drawbacks of this technology. Recent information fromPhillips indicates that this process is being adapted fordesulfurization of diesel.

Various other desulfurization processes are known or have been proposed.For example, U.S. Pat. No. 3,063,936, issued on Nov. 13, 1962 to Pearceet al. discloses that sulfur reduction can be achieved for straight-runnaphtha feedstocks from 357 ppmw to 10-26 ppmw levels by hydrotreatingat 380° C. using an alumina-supported cobalt molybdate catalyst.According to Pearce et al., a similar degree of desulfurization may beachieved by passing the straight-run naphtha with or without hydrogen,over a contact material comprising zinc oxide, manganese oxide, or ironoxide at 350 to 450° C. Pearce et al. propose to increase sulfur removalby treating the straight run naphtha feeds in a three-stage process inwhich the hydrocarbon oil is treated with sulfuric acid in the firststep, a hydrotreating process employing an alumina-supported cobaltmolybdate catalyst is used in the second step, and an adsorptionprocess, preferably using zinc oxide is used for removal of hydrogensulfide formed in the hydrotreating step as the third step. The processis said to be suitable only for treating feedstocks that aresubstantially free from ethylenically or acetylenically unsaturatedcompounds. In particular, Pearce et al. disclose that the process is notsuitable for treating feedstocks, such as hydrocarbons obtained as aresult of thermal cracking processes that contain substantial amounts ofethylenically or acetylenically unsaturated compounds such as full-rangeFCC naphtha, which contains about 30% olefins.

U.S. Pat. No. 5,157,201 discloses that organic sulfur species, primarilycomprising organic sulfides, disulfides, and mercaptans, can be adsorbedfrom olefin streams, without saturating the olefins, by contacting thefeed with a metal oxide adsorbent at relatively low temperatures (50 to75° C.), in the absence of hydrogen. The metal oxide adsorbent includesmetal oxides selected from a group consisting of a mixture of cobalt andmolybdenum oxides, a mixture of nickel and molybdenum oxides and nickeloxide supported on an inert support. The adsorbed organic sulfurcompounds are removed from the sorbent by purging with an inert gaswhile heating at a temperature of about 200° C. for at least about 45minutes. Although such low-temperature adsorption processes avoid anyolefin saturation, these processes are limited to removal of lightersulfur compounds such as mercaptans and organic sulfides and disulfides.These processes cannot be used effectively for removal of thiophenes,benzothiophenes, and higher cyclic sulfur compounds, which typicallyaccount for greater than 50% of the sulfur in FCC naphtha.

In summary, currently available and proposed technologies for reducingsulfur content of FCC naphtha feedstocks to levels of 30 ppmw or lessare capital intensive, operationally complex, typically requiresignificant hydrogen consumption, can severely reduce octane numbervalues and/or result in loss in yield, and rely on expensivehydrotreating catalysts in whole or in part. In addition, the existingand proposed technologies rely on fixed-bed or bubbling-bed reactorsresulting in limited throughputs and substantial capital investment.

SUMMARY OF THE INVENTION

The present invention accomplishes sulfur reduction in gasoline anddiesel fuels, components and precursors of gasoline and diesel fuelssuch as naphthas, i.e., full and medium range FCC naphthas, cokernaphthas, straight run naphthas, visbreaker naphthas, and thermallycracked naphthas, light cycle oils, coker distillates, straight-rundiesel, hydrocracker diesel, and the like, without relying onhydrotreating processes that employ costly transition metal HDScatalysts. Accordingly, the invention can minimize or eliminate variousknown disadvantages of conventional and proposed desulfurizationprocesses for producing low-sulfur gasoline and diesel fuels, includingoctane number loss, olefin content reduction, and/or yield loss indesulfurized products, hydrogen consumption and its associated costs,the high cost of manufacturing and regenerating HDS catalysts, and thedisposal costs associated with various environmentally undesirable HDScatalysts. In preferred embodiments, the present invention canaccomplish substantial sulfur removal at high throughput levels, therebyallowing a significant reduction in the capital investment required toachieve large scale production of low-sulfur gasoline, diesel, andrelated fuels.

In accordance with one aspect of the present invention, a normallyliquid hydrocarbon fuel or fuel component, such as an FCC naphtha, FCClight cycle oil, coker distillate, straight run diesel fraction, or thelike, is treated at an elevated temperature, preferably a temperatureabove about 300° C. (572° F.), with an active metal oxide sulfursorbent, preferably a zinc oxide-based or iron oxide-based sorbent, inthe absence of an active HDS catalyst, to reduce sulfur contaminantlevels to less than about 30 ppmw, sulfur. Sulfur-laden sorbent isseparated from the desulfurized hydrocarbon product and is preferablyregenerated by treatment with an oxygen-containing gas, e.g., air, andthen recycled for use in the desulfurization operation. The invention isapplicable to hydrocarbon fuels and to hydrocarbon fuel fractions andprecursors, of various sulfur contents, for example: FCC naphtha havingan average sulfur content of between about 150 and about 3,000 ppmw,more typically, between about 500 to about 2,000 ppmw; diesel fuelblends, precursors and fractions such as light cycle oil, cokerdistillate and straight run diesel fractions having an average sulfurcontent between about 5,000 and about 30,000 ppmw, more typically,between about 7,000 and about 20,000 ppmw. The process of this inventionis equally applicable to partially desulfurized feedstocks such ashydrotreated FCC naphtha and diesel, to reduce their sulfur content tobelow 30 ppmw.

The process of the invention can be carried out with or without additionof hydrogen to the feed; however, it is preferred to add a sufficientamount of hydrogen to the feed to avoid coking of the feed as it isheated to the elevated temperatures required for desulfurization.Because no active HDS catalyst is used in the present process, hydrogenaddition to minimize coking can typically be achieved with minimal orsubstantially no hydrogen consumption so that the hydrogen can berecovered from the desulfurized process effluent and recycled. Moreover,because of the substantial absence of an HDS catalyst, saturation ofdesirable olefins in the hydrocarbon feed can be avoided or minimizedeven at high temperature reaction conditions, and even in the presenceof added hydrogen. Furthermore, the hydrogen gas stream used in theprocess can be of relatively low purity; for example, a waste streamcontaining hydrogen, as may be found in a refinery or petrochemicalplant. Moreover, because no active HDS catalyst is required in thepresent invention, no hydrogen treatment is required for regeneration orreactivation of the sorbent.

The present inventors have further found that the active metal oxidesulfur sorbents, particularly zinc oxide-based and iron oxide-basedsorbents, when used in combination with a refractory inorganic oxidecracking catalyst, e.g., alumina, are capable of removing both straightchain organic sulfur components such as organic sulfides, disulfides,and mercaptans, and cyclic organic sulfur components includingsubstituted and unsubstituted thiophenes, benzothiophenes, and, to someextent, dibenzothiophenes from hydrocarbon fuels, their fractions andprecursors, without hydrotreating. In this regard, the present inventorshave discovered that a refractory inorganic oxide cracking catalyst,such as alumina, silica, an aluminosilicate or a metal stabilizedrefractory inorganic oxide cracking catalyst such as metal stabilizedalumina, when used to support, or otherwise in combination with theactive metal oxide sulfur sorbent, has catalytic activity forselectively cracking cyclic organic sulfur compounds to provide ahydrocarbon and a sulfur species. The sulfur species can be captured bythe cracking catalyst or by the active metal oxide sulfur sorbent as ametal sulfide or a metal-sulfur complex. Although prior art processeshave primarily relied on hydrotreating of FCC naphthas and diesel fuelfractions and components using HDS catalysts to convert organic sulfurcontaminants to hydrogen sulfide, followed by amine and Claus processtreatments for removal of hydrogen sulfide, it has now been found thatactive metal oxide sorbents, preferably zinc oxide-based and ironoxide-based sorbents, supported on or otherwise combined with arefractory inorganic oxide cracking catalyst, can directly removeorganic sulfur contaminants from hydrocarbon feedstocks at elevatedtemperatures without requiring use of an active HDS catalyst. In turn,detrimental aspects of hydrotreating-desulfurization processes, such asoctane number reduction, and/or olefins loss, can be minimized oravoided in accord with the present invention.

The active metal oxide sulfur sorbents and refractory inorganic oxidecracking catalyst are preferably used simultaneously to treat thehydrocarbon fuel feed; however they can alternatively be usedsequentially in the process of the invention. In preferred embodimentsin which the active metal oxide sulfur sorbent and the refractoryinorganic oxide cracking catalyst are used simultaneously, the activemetal oxide sulfur sorbent is supported on or combined with a refractoryinorganic oxide cracking catalyst such as alumina, silica,aluminosilicate, zeolite or the like. This can also provide hightemperature stability and extremely high attrition resistance to thesorbent particles.

According to another aspect of the invention, it has been found thatcertain metal-substituted refractory inorganic oxide cracking catalystscan remove organic sulfur compounds from hydrocarbon feeds, and can alsoremove sulfur from at least some of the organic sulfur compounds inhydrocarbon feeds, particularly cyclic sulfur compounds such asthiophenes and benzothiophenes, without requiring use of an HDS catalystor hydrotreating of the feed. The metal, which can be zinc in onecurrently preferred embodiment, or iron in another currently preferredembodiment, is more generally selected from the group of metals, whichare capable in their oxide form, of removing reduced sulfur compoundsfrom gaseous streams by conversion of the metal oxide to a metalsulfide, such metal oxides being known in the art. The refractoryinorganic oxide cracking catalyst can be fully, or only partially,reacted with the metal. The metal-substituted refractory inorganic oxidecracking catalyst can be prepared according to processes well known inthe art and is advantageously prepared by partially or fully reacting ametal oxide sulfur sorbent with a refractory inorganic oxide crackingcatalyst, such as alumina, silica, an aluminosilicate or the like, toform the corresponding metal aluminate, silicate, aluminosilicate or thelike. Suitable active metal oxide sorbents for use in the process of theinvention include sorbents based on zinc oxide, zinc titanate, zincferrite, iron oxide, iron titanate, manganese oxide, cerium oxide,copper oxide, copper cerium oxide, copper ferrite, copper titanate,copper chromium oxide, vanadium oxide, calcium oxide, calcium carbonate,magnesium oxide, magnesium carbonate, and mixtures thereof.

In particular, the metal-substituted inorganic oxide cracking catalystsorbent, i.e., metal aluminate, silicate, aluminosilicate or the like,can achieve full or partial conversion of organic sulfur compounds,including cyclic sulfur compounds such as thiophenes andbenzothiophenes, to a metal sulfide or a metal-sulfur complex. Suchmetal-substituted inorganic oxide cracking catalyst sorbents can be usedin accordance with the invention to treat a hydrocarbon fuel component,precursor, or blend, preferably an FCC naphtha, or a diesel fuelprecursor, component, or blend, at an elevated temperature, preferablyabove about 300° C. (572° F.), and the treated hydrocarbon stream isthen separated from the sulfur-laden sorbent to provide a hydrocarbonproduct having a sulfur contaminant level preferably of less than about30 ppmw, without requiring hydrotreating of the feed using an active HDScatalyst. Moreover, such metal-substituted inorganic oxide crackingcatalyst sorbents also possess high mechanical strength and attritionresistance. Currently preferred metal-substituted inorganic oxidematerials include zinc aluminate, iron aluminate and combinationsthereof.

In preferred embodiments of the invention, the sulfur-laden sorbentemployed in the desulfurization process of the invention is regenerableby treatment with oxygen at an elevated temperature. According to onecurrently preferred embodiment of the invention, the regenerable sorbentis an active metal oxide sulfur sorbent supported on, or otherwisecombined with a metal-substituted refractory inorganic oxide crackingcatalyst, wherein all or a portion of the metal component of themetal-substituted refractory inorganic oxide is the same metal as themetal of the active metal oxide sulfur sorbent. In particular, suchregenerable sorbents are used to remove sulfur compounds from ahydrocarbon fuel component feed, to achieve sulfur contaminant levels ofless than about 30 ppmw of total sulfur in the product effluent, withoutrequiring hydrotreating of the feed using an active HDS catalyst. Thecombination of the metal oxide sulfur sorbent and metal refractoryinorganic oxide cracking catalyst, e.g., zinc oxide/zinc aluminate oriron oxide/iron aluminate, can be particularly desirable to prevent orminimize deactivation of the sulfur removal activity of the sorbentduring the adsorption-regeneration process. In a currently preferredembodiment, a zinc titanate and/or iron oxide sorbent is supported on analumina or a metal aluminate, preferably zinc and/or iron aluminate,support.

The sulfur-laden sorbent used to remove sulfur compounds fromhydrocarbon feedstocks in the process of the present invention, isregenerated by contacting the sorbent with an oxygen-containing gas,preferably air, at a temperature sufficient to cause the sulfur presenton the sorbent to react with oxygen to form sulfur dioxide. Typically,the equilibrium temperature in the regeneration zone will exceed atemperature of about 425° C. (800° F.). In one preferred embodiment ofthe invention, regeneration can be initiated or supplemented by additionof the metal sulfide additives disclosed in U.S. Pat. No. 5,914,288,issued on Jun. 22, 1999 to Turk et al.; the disclosure of which isincorporated herein by reference. As disclosed in the aforesaid Turk etal. patent, a preferred metal sulfide initiator is iron pyrite mineralore.

The regeneration reaction converts the sulfur-laden sorbent, to theactive metal oxide form, for example, to zinc or iron oxide, zinctitanate, or zinc or iron aluminate, and the regenerated sorbent isreturned directly to the desulfurization zone. Because the sorbents usedin the process of the present invention do not include an active HDScatalyst component, no separate hydrogenation treatment is necessary forregenerating the sorbents to an active state. Accordingly, the energycost, hydrogen consumption, and reaction vessels required for hydrogentreatment of hydrogenation catalysts are avoided in the process of thepresent invention.

In one preferred embodiment the invention, the desulfurization processis carried out employing a transport bed reactor with a vapor residencetime of less than about 20 seconds, more typically less than about 10seconds. Nevertheless, high sulfur containing hydrocarbon feedstocks,i.e., having a sulfur content greater than about 150-300 ppmw, moretypically greater than about 600 ppmw, can be desulfurized in accordwith the invention to achieve sulfur reduction to less than 30 ppmw,more typically less than 10 ppmw. The extremely high throughput processaccording to this aspect of the invention greatly reduces capitalinvestment since a relatively small reactor can be used for treatingsubstantial quantities of hydrocarbon feedstocks. Use of a highthroughput transport reactor is possible because of the extremely highattrition resistance of preferred sorbents used in the presentinvention. This unique combination of extremely high attritionresistance, allowing these sorbents to be used in a transport reactor,and relatively high activity for selectively cracking cyclic sulfurcompounds in hydrocarbon feedstocks combined with sorption activity ofactive metal oxide component of the sorbent for various inorganic andorganic sulfur compounds provides significant benefits and advantages ascompared to processes of the prior art.

In another preferred embodiment of the invention, the desulfurizationprocess is carried out employing a bubbling bed reactor to treathydrocarbon fuel feedstocks having an initial sulfur content greaterthan about 150-300 ppmw, more typically greater than about 600 ppmw, inorder to achieve sulfur reduction to less than 30 ppmw, more typicallyless than 10 ppmw. Bubbling bed reactors, which can provide excellentgas-solid contact and significant process and capital cost benefits ascompared to prior art fixed and packed bed processes, can be employed inaccord with the invention using various preferred, high attritionresistance sorbents.

According to another aspect of the invention, sulfur contaminants areremoved from an FCC hydrocarbon stream by treating the stream underconventional FCC process conditions, with a regenerable sorbentcomprising an active metal oxide sulfur sorbent supported on, orotherwise combined with a refractory inorganic oxide cracking catalyst,preferably comprising a metal substituent, as discussed previously.Advantageously, desulfurization of the FCC hydrocarbon process stream isaccomplished simultaneously with the FCC process by adding the sorbentto the FCC riser, e.g., as an additive to the FCC catalyst. According tothis aspect of the invention, sulfur compounds initially present in theFCC feedstock, or generated during the FCC process, are selectivelycaptured by the sorbent in the FCC riser. The sulfur-laden sorbent isthen sent to the FCC regenerator along with the carbon-laden FCCcatalyst where it is regenerated by the oxygen-containing gas, typicallyair, which is used to regenerate the FCC catalyst. During regeneration,sulfur carried by the sorbent is converted to a sulfurdioxide-containing gas stream that can be treated for sulfur removal ina downstream process unit such as a sulfur dioxide scrubber.

Desulfurization in combination with an FCC operation according to thisaspect of the invention is particularly desirable since most of thesulfur (>90%) in gasoline comes from the naphtha produced byconventional FCC treatment. In this regard, the FCC operation is used toupgrade the less desirable portions in crude oil as is well known tothose skilled in the art. Because such less desirable portions of oilinclude substantial quantities of undesirable sulfur-containingcomponents, the product streams generated by the FCC unit also have highsulfur contents. Thus, although some of the sulfur initially in the feedto a conventional FCC unit is removed as H₂S generated during crackingand is collected as non-condensable gas, a substantial portion of thesulfur remains in the FCC product as organic sulfur contaminants,distributed among the various FCC product fractions including FCCnaphtha, light cycle oil (LCO), heavy cycle oil (HCO) and the bottomsfraction. Typical sulfur compounds found in FCC naphtha and LCO areessentially heavy thiophenic materials, which are very difficult toconvert into H₂S during the catalytic cracking process in a FCC reactor.

According to this aspect of the invention, the active metal oxide sulfursorbent is added to the FCC catalyst in an amount sufficient to achieveremoval of at least about 50 wt. % of sulfur compounds from the FCCnaphtha product, i.e., the FCC liquid product fraction having a finalboiling point (FBP) less than about 430° F. More preferably, the activemetal oxide sulfur sorbent is also active for removal of sulfurcontaminants from heavier FCC product fractions and is added to the FCCcatalyst in an amount sufficient to achieve removal of at least about 50wt. % of sulfur compounds from the FCC naphtha and LCO productfractions, i.e., the FCC liquid product fraction having an FBP of lessthan about 650° F. In currently preferred embodiments according thisaspect of the invention, the active metal oxide sulfur sorbent is addedto the FCC catalyst in an amount of from about 1 to about 10 wt %, basedon the weight of the FCC catalyst.

BRIEF DESCRIPTION OF THE DRAWINGS

In the drawings which form a portion of the original disclosure of thisapplication:

FIG. 1 is a schematic view of a preferred desulfurization andregeneration process according to the present invention; and

FIG. 2 is a schematic view illustrating an FCC desulfurization processin accordance with another preferred aspect of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

The present invention now will be described more fully hereinafter withreference to the accompanying drawings, in which preferred embodimentsof the invention are shown. This invention may, however, be embodied inmany different forms and should not be construed as limited to theembodiments set forth herein; rather, these embodiments are provided sothat this disclosure will be thorough and complete, and will fullyconvey the scope of the invention to those skilled in the art. Likenumbers refer to like elements throughout.

FIG. 1 illustrates a preferred hydrocarbon feedstock desulfurizationprocess according to the present invention. As shown in FIG. 1, theprocess includes a desulfurization zone 10 and a regeneration zone 20.In a preferred process according to the invention, and illustrated inthe drawing, each of the desulfurization zone 10, and the regenerationzone 20, is defined by a transport bed reactor. It will be apparent tothe skilled artisan however that other conventional fluidized bedreactors, including bubbling bed, circulating bed, and riser reactorscan be used in the process of the invention. In addition, thehydrocarbon feedstock desulfurization process of the present inventioncan be conducted using other conventional catalytic reactors includingfixed bed and moving bed reactors, such reactors being well known tothose skilled in the art.

Preferred transport bed reactors are similarly known to those skilled inthe art and are described in, for example, Campbell, William N. andHenningsen, Gunnar B., Hot Gas Desulfurization Using Transport Reactors,publication from the M. W. Kellogg Company, pp 1059-64, 12th AnnualInternational Pittsburgh Coal Conference Proceedings, September 1995,which is incorporated in its entirety herein by reference. Transport bedreactors are also described in U.S. Pat. No. 5,447,702, issued on Sep.5, 1995 to Campbell et al., which is incorporated herein in its entiretyby reference.

As illustrated in FIG. 1, a vaporized sulfur containing hydrocarbonfeedstock 30, which can be FCC naphtha, is fed at a predeterminedvelocity through an inlet 32 into the desulfurization zone 10 inadmixture with a sulfur sorbent comprising an active metal oxidesorbent, or a metal-substituted refractory inorganic oxide crackingcatalyst, preferably a sorbent comprising both, i.e., an active metaloxide sorbent supported on, or otherwise combined with ametal-substituted refractory inorganic oxide cracking catalyst. Thehydrocarbon feed 30, including added sorbent, is fed by means of inlet34 at a temperature between about 300° C. (572° F.) and about 600° C.(1112° F.), preferably at a temperature between about 371° C. (700° F.)and about 538° C. (1000° F.). Optional hydrogen feed 36 is alsointroduced into the desulfurization zone 10 via inlet 32. The combinedhydrogen, hydrocarbon and sorbent stream is transported upwardly througha riser pipe 38 during a relatively short time period of less than about20 seconds, typically less than about 10 seconds for achievingdesulfurization of the feed stream 30. Typically, the superficial gasvelocity is between about 5 and about 40 ft/sec, more preferably betweenabout 10 and about 30 ft/sec. The desulfurization zone 10 may have morethan one section. In one of the preferred option, the desulfurizationzone 10 will consist of two sections, namely a mixing zone in the bottomand a riser zone at the top. The relative length and diameter of thesesections will depend on the kinetics of desulfurization reaction,residence time required, sulfur content of the hydrocarbon feedstock,and feedstock throughput, as will be well known to those skilled in theart.

The hydrocarbon feedstock 30 treated in accordance with the process ofthe present invention is preferably a normally liquid hydrocarbon fuelor fuel component. The term “normally liquid” means liquid at StandardTemperature and Pressure (STP) conditions as will be apparent to theskilled artisan. Although the feedstock 30 is an FCC naphthaconstituting a component or fraction of an automotive gasoline fuel inone preferred embodiment of the invention, the invention is equallyapplicable to other hydrocarbon fuel feedstocks, and to precursors andcomponents thereof. In particular, the invention is applicable to dieselfuel, aviation fuel, and the like, and to components and precursorsthereof including, for example, coker naphthas, thermally crackednaphthas, full-range FCC naphthas, light cycle oils, straight-rundistillate fractions, and the like. In this regard, it will beappreciated that the hydrocarbon feedstock 30 supplied to thedesulfurization zone 10, can have differing boiling point ranges, andwill contain varying levels of various organic sulfur contaminantstypically including organic sulfides and disulfides, mercaptans,substituted and unsubstituted thiophenes, benzothiophenes, anddibenzothiophenes. In the case of FCC naphtha, the concentration ofthese sulfur compounds depends on boiling point cut from thefractionator and sulfur content of the feed to the FCC, and typicallyexceeds 150 ppmw, and more typically exceeds 300 ppmw as discussedpreviously. In the case of diesel fuel components and blends, the sulfurcontent is typically higher. In particular, diesel is typically formedfrom a blend comprising light cycle oil recovered from an FCC unit, adistillate recovered from a coker unit (coker distillate), and astraight-run fraction recovered from the crude fractionation unit. Lightcycle oils and coker distillates typically have sulfur contents in therange of from about. 5,000 to about 30,000 ppmw. Straight-run fractionsused in diesel fuels can be derived from sweet or sour crude, andtypically have different sulfur content ranges, which in the case ofsweet crude straight-run fractions, range from about 300 to about 5,000ppmw, and in the case of sour crude straight-run fractions, range fromabout 5000 to about 30,000 ppmw. In turn, the complete diesel fuelblend, prior to a conventional hydrotreating step, typically has asulfur content of up to about 2000 ppmw, and in some cases can have asulfur content ranging from about 5000 to about 30,000 ppmw.

The process of the invention is equally applicable to achievesubstantial sulfur reduction in partially desulfurized feedstocks suchas hydrotreated FCC naphtha and hydrotreated diesel blends andcomponents to reduce their sulfur content to below 30 ppmw, whileavoiding olefin saturation, product yield losses and/or increasedprocessing costs which can accompany sulfur removal by HDS processes,particularly in the case of cyclic and polycylic organic sulfurcontaminants. In particular, the desulfurization process of theinvention can be employed to accomplish a polishing step or the like forremoval of cyclic and polycylic organic sulfur contaminants fromrelatively low-sulfur feedstocks, in order to achieve removal of atleast about 25 wt. %, more preferably at least about 50 wt. %, of thecyclic and polycyclic organic sulfur contaminants initially present in alow-sulfur hydrocarbon fuel, fuel component or fuel precursor feed.

In embodiments of the invention wherein diesel fuels and/or theircomponents or precursors are treated to reduce sulfur, the preferredprocess conditions and/apparatus can accordingly be varied depending onthe particular feedstock, and sulfur content as will be apparent tothose of skill in the art. Thus, when a diesel fuel, or precursor(s) orcomponent(s) thereof, is treated for sulfur removal in the processillustrated in FIG. 1, a high sulfur diesel feed 30, is fed in vaporform into the desulfurization zone 10 in admixture with an active metaloxide sorbent at a temperature of between about 350° C. (662° F.) andabout 750° C. (1382° F.), preferably at a temperature between about 450°C. (842° F.) and about 700° C. (1292° F.). The combined diesel feed andsorbent stream, with or without optional hydrogen feed 36 is transportedupwardly through riser pipe 38 during a relatively short residence timeof less than about 20 seconds, to thereby achieve desulfurization of thediesel feed 30.

Although not specifically illustrated in the drawings, thedesulfurization process of the invention can be advantageously carriedout employing a conventional bubbling bed reactor to accomplishgas-solid contact between the hydrocarbon fuel feedstock and the activemetal oxide sorbent. Bubbling bed reactors can be advantageouslyemployed to treat any of the various fuels, fuel components, and fuelprecursors discussed previously, and can be particularly beneficial fortreating hydrocarbon fuels and fractions having boiling point rangesexceeding that of FCC naphtha in view of the enhanced gas-solid contactthat can be achieved in bubbling bed reactors as compared to transportbed reactors. Bubbling bed reactors provide excellent gas-solid contactand significant process and capital cost benefits as compared to fixedand packed bed reactors which are typically used in prior arthydrodesulfurization processes in order to minimize olefin saturationand product yield losses. The active metal oxide sulfur sorbent employedto treat hydrocarbon feedstocks in bubbling bed reactors according tothis embodiment of the invention, is advantageously a high attritionresistance sorbent, discussed in greater detail below. As indicatedpreviously, the desulfurization process of the present invention canalternatively be conducted using other conventional catalytic reactorsincluding fixed bed and moving bed reactors with substantial benefits ascompared to prior art hydrodesulfurization processes.

The active metal oxide sulfur sorbent employed in the invention includesat least one active metal oxide capable of removing sulfur compoundsfrom the sulfur-containing fuel feed stream to form a metal sulfide or ametal-sulfur complex. The term “active metal oxide sulfur sorbent” asused herein refers to active metal oxides and mixed active metal oxides,including different oxides of the same elements, for example, zinctitanate which includes various oxides of the formula ZnO.n(TiO₂), orvarious iron oxides of the formula Fe_(x)(O)_(y), and to mixed oxides ofdifferent metals including active metal oxides derived from calcining ofactive metal oxides, and also to carbonates. Such active metal oxidesorbents can include binders that are mixed or reacted with the activemetal oxide, supports that support the metal oxide, and the like as willbe apparent to the skilled artisan. Advantageously, the sorbents used inthe present invention are regenerable by treatment with oxygen at anelevated temperature. For purposes of the present invention, a sorbentis considered regenerable when it can be used for desulfurization of ahydrocarbon feed, and can thereafter be reactivated at least once bytreatment with oxygen at an elevated temperature, to a sulfur removalactivity level greater than 50% of the original sulfur activity level ofthe sorbent (based on the original weight percent sulfur adsorbingcapacity of the sorbent under the same conditions). Active metal oxidesorbents exhibiting good adsorption rates and capacity for sulfurcompounds, good regenerability without appreciable loss of efficiency orefficacy, and high attrition resistance are preferred for use in thisinvention. These sorbents chemically react with the sulfur atoms of theorganic sulfur compounds in the feed stream and the active metal oxideis thus converted into a metal sulfide and/or a metal-sulfur complex.

Suitable active metal oxide sorbents for use in the process of theinvention, include sorbents based on zinc oxide, zinc titanate, zincaluminate, zinc silicate, zinc ferrite, iron oxide, iron aluminate, ironzinc oxide, manganese oxide, cerium oxide, copper oxide, copper ceriumoxide, copper titanate, copper chromium oxide, copper aluminate,vanadium oxide, calcium oxide, calcium carbonate, magnesium oxide,magnesium carbonate, and mixtures thereof, particularly mixtures of zincoxides with an iron oxide, and/or copper oxide.

In one particularly preferred embodiment of the invention, the activemetal oxide is supported on or otherwise combined with a refractoryinorganic oxide cracking catalyst support. Refractory inorganic oxidecracking catalyst support materials are well known to those skilled inthe art and include various aluminas, silicas, aluminosilicates, andzeolites. Refractory inorganic oxide cracking catalysts supportmaterials which have been reacted with a metal or metal oxide, such asmetal or metal oxide aluminates, metal or metal oxide silicates, metalor metal oxide aluminosilicates, and metal or metal oxide zeolites arecurrently preferred for use in the present invention. One particularlypreferred supported active metal oxide for use in the present inventionis a zinc aluminate supported zinc titanate as disclosed in PCTApplication WO 99/42201 A1, published Aug. 26, 1999, entitled “AttritionResistant, Zinc Titanate-Containing, Reduced Sulfur Sorbents”, which ishereby incorporated herein by reference. Other metal oxide aluminatesupports described in the aforesaid PCT Application are also suitablefor use in the present invention. The metal oxide aluminate supportedzinc titanate sorbent materials can be formulated to be highly attritionresistant even at high temperatures, while maintaining substantialchemical activity and regenerability. Other metal and metal oxidealuminates such as iron aluminates, and/or copper aluminates, are also,or alternatively, desirably employed in preferred embodiments of theinvention to likewise provide high attrition resistance along withsubstantial sulfur-removal capacity and good regenerability.

Although the active metal oxide sulfur sorbent is preferably supportedby, or combined with, the refractory inorganic oxide cracking catalystso that the hydrocarbon fuel stream is treated simultaneously by theactive metal oxide sorbent and the refractory inorganic oxide crackingcatalyst, the present invention also includes processes in which thehydrocarbon fuel stream is treated with the refractory inorganic oxidecracking catalyst and the active metal oxide sorbent sequentially, forexample, by passing the hydrocarbon fuel stream through sequentialtreatment zones including the respective refractory inorganic oxidecracking catalyst and metal oxide sorbent.

Mixed active metal oxide sulfur sorbents are particularly desirable insome advantageous embodiments of the invention. For example, it is knownthat the sulfur adsorption capabilities of active metal oxide sorbentsvary from sorbent to sorbent at different temperatures. It has beenfound that the reaction kinetics associated with sulfur conversion andsorption by zinc oxide-based sorbents can be substantially enhanced attemperatures below about 525° C. (1000° F.) by incorporating a minoramount of an active metal sorbent which adsorbs sulfur at lowertemperatures than zinc oxide sorbents. One such preferred additionalactive metal oxide sorbent is copper oxide which may be included in anamount ranging from about 5 to about 45 weight percent, preferably about5 to about 20 weight percent based on the weight of the active zincoxide component (for example, zinc titanate). Other promoters mayinclude oxides of iron, silver, gold, or any combination thereof. Otherdesirable mixed metal oxide sorbents include iron oxides mixed with zincoxides and/or zinc titanates and/or copper oxides.

Numerous other active metal oxide sorbents can also be used in theprocess of the invention. Exemplary active metal oxide sorbents aredisclosed in U.S. Pat. No. 5,254,516, issued Oct. 19, 1993 to Gupta etal., U.S. Pat. No. 5,714,431, issued Feb. 3, 1998 to Gupta et al., andU.S. Pat. No. 5,972,835, issued Oct. 26, 1999 to Gupta. Still otherexemplary active metal oxide sorbents include sorbents which aremarketed by Philips Petroleum Company and contain a zinc oxide-basedsorbent (but without any substantial nickel or any other Group 6, 8, 9,or 10 metal other than iron). Other useful metal oxide sorbent materialsinclude those disclosed in U.S. Pat. Nos. 5,866,503, 5,703,003, and5,494,880, issued Feb. 2, 1999, Dec. 30, 1997, and Feb. 27, 1996,respectively, to Siriwardane. The latter are commercially available asRVS materials from SudChemie Inc.

Returning to FIG. 1, the sorbents fed into the desulfurization zone 10via inlet pipe 34 are preferably substantially free from activehydrodesulfurization catalysts. The term “active hydrodesulfurizationcatalyst(s)” is used herein to mean nickel, cobalt, molybdenum,tungsten, and combinations of these metals when present in a state thatis chemically active or activatable for hydrodesulfurization. Suchmetals are considered active or activatable for hydrodesulfurization, ina sulfided state, or in a form that is readily converted to the sulfidedmetal when exposed to a hydrocarbon feed containing hydrogen and sulfurcontaminants at high temperature desulfurizing conditions. Inparticular, sulfides of nickel, cobalt, molybdenum, tungsten andcombinations thereof, are well known by those skilled in the art to bethe active catalytic components for hydrodesulfurization. It is likewisewell known in the art that oxides of molybdenum, cobalt, nickel, andtungsten can be readily converted to the active sulfides by exposure tohydrogen and sulfur compounds in hydrocarbon feeds at thedesulfurization conditions employed in this invention.

Each of the terms, “substantially free” and “substantial absence”, asapplied to active hydrodesulfurization catalysts, is used herein to meanthat active hydrodesulfurization catalyst(s) are not present, in a formphysically accessible to the hydrocarbon feed and in sufficientquantity, to promote substantial conversion of the organic sulfurcomponents in the feedstock into H₂S by reaction with hydrogen gas,under the desulfurization conditions employed in a process of theinvention. In turn, saturation of desirable hydrocarbon olefins in thefeed is substantially reduced or eliminated, even in the presence ofsmall quantities of hydrogen, and even at high temperatures. Similarlythe costs associated with hydrogen consumption can be greatly reduced orsubstantially eliminated.

Preferably, the sorbents used in the present invention contain less thanabout 1.0 wt. % nickel, cobalt, molybdenum, tungsten and/or combinationsof these metals, calculated based on the weight of such metal(s), and onthe total sorbent weight including the cracking catalyst support orcomponent. More preferably, the sorbents used in the present inventioncontain less than about 0.5 wt. % nickel, cobalt, molybdenum, tungstenand/or combinations of these metals, calculated based on the weight ofsuch metal(s), and on the total sorbent weight. Even more preferably thesorbents used in the present invention contain less than about 1.0 wt. %of Group 6 and/or Group 8, 9, and 10 metals (excluding iron), and mostpreferably the sorbents used in the present invention contain less thanabout 0.5 wt. % of Group 6 and/or Group 8, 9, and 10 metals (excludingiron), calculated based on the weight of such metal(s), and on the totalsorbent weight including the cracking catalyst support or component.

Returning to FIG. 1, the sorbent added via inlet pipe 34 is transportedupwardly through riser pipe 38 and separated via a cyclone separator 42.The separated sorbent is recovered via a standpipe 44 and a portion ofthe sorbent is passed via a pipe 46 to the regeneration zone 20 whichpreferably constitutes a riser pipe 50. An oxygen-containingregeneration gas 52, which is preferably ambient air, is added to theriser 50 via inlet pipe 54. In addition, fresh makeup sorbent 56 isadded as necessary via inlet pipe 54. Further, the metal sulfideadditives for enhancing or initiating regeneration, described in theaforementioned Turk et al. U.S. patent, can be advantageously added tothe riser 50 via line 58 and inlet 54 in order to improve processeconomies in the regeneration zone 20 as described in greater detail inthe aforementioned Turk et al. patent.

Preferably, the heat carried by the heated sorbent particles admitted tothe riser 50 via pipe 46, and the heat carried by the oxygen in theoxygen-containing stream, are sufficient to establish conditions in theregeneration zone 20 for initiating regeneration of the sulfided activemetal oxide sorbent and/or for initiating reaction of the metal sulfideadditive, added via line 58, with oxygen in a highly exothermiccombustion reaction to form a metal oxide and sulfur dioxide. The heatreleased by the metal sulfide additive can, in some cases, be used toinitiate regeneration of the active metal oxide sulfur adsorbent atstart-up of the process, or can be used as a supplemental heating sourcefor maintaining the desired temperature in the regeneration zone 20.

The temperature in the regeneration zone during the regenerationreaction typically is within a range of from about the same temperatureas the temperature in the desulfurization zone 10 up to a temperature ofabout 200° C. higher than the temperature in zone 10, for example, atemperature of about 425° C. (800° F.) or higher under steady stateconditions. The heat generated during removal of the sulfidecontaminants from the active metal oxide sorbents advantageouslysupplies all or a portion of the heat necessary for vaporization of thehot feed gas stream 30.

In the regeneration zone 20, the oxygen containing regeneration gasreacts with the sulfur on the active metal oxide sorbent to producesulfur oxides which are removed as a tail gas stream via line 60.Regenerated sorbent is separated via a cyclone separator 62 and passedvia a standpipe 64 and inlet pipe 34 back to the desulfurization zone10.

A desulfurized hydrocarbon fuel stream 70 is recovered from cycloneseparator 42 and passed to a conventional separation zone 72 forseparation of a recycle hydrogen stream 74 and a desulfurizedhydrocarbon fuel stream 76.

The desulfurization process of the present invention can be used totreat naphtha and diesel streams having sulfur contents of from 150 ppmwto over 3,000 ppmw, while reducing the sulfur contaminants by virtuallyany pre-selected amount. As will be apparent to those skilled in theart, the percentage of sulfur reduction can be readily controlled byvarying residence time and temperature in the desulfurization zone.

Advantageously, the process of the invention is conducted at conditionsresulting in a sulfur content reduction of at least about 50% or more,preferably at least 80%, more preferably at least about 90%, even morepreferably at least about 95%, based on the sulfur content, by weight,of the feedstock. In preferred embodiments of the invention, the sulfurcontaminants can be reduced to levels below 20 ppmw, more preferablybelow 10 ppmw during a residence time preferably below about 20 seconds,more preferably below about 10 seconds. Moreover, such sulfur reductionsare preferably achieved with an octane number loss, in the case of FCCnaphtha of less than about 5, preferably less than about 2.

With reference now to FIG. 2, an FCC desulfurization process inaccordance with another preferred aspect of the present invention isillustrated by a schematic view wherein certain of the drawing parts arelabeled with the same numbers as in FIG. 1, and accordingly representthe same parts as the corresponding parts numbered the same in FIG. 1.

In particular, FIG. 2 illustrates a preferred process of the inventionin which sulfur contaminants are removed from an vaporizedsulfur-containing FCC feedstock 130 simultaneously with an otherwiseconventional FCC process which is conducted in a conventional FCC riserreactor 110 under conventional temperature, pressure and residence timesemployed for FCC processes. A mixture of a conventional FCC catalystwith a regenerable sorbent comprising an active metal oxide sulfursorbent supported on, or otherwise combined with a refractory inorganicoxide cracking catalyst, preferably comprising a metal substituent, isfed to the FCC reactor zone 110 via line 140. Although not specificallyshown in FIG. 2, the FCC catalyst and the regenerable sorbentalternatively can be admitted to the FCC riser 138 via separate lines,or by mixing with the vaporized sulfur-containing FCC feedstock 130.According to this aspect of the invention, sulfur compounds initiallypresent in the FCC feedstock, or generated during the FCC process, areselectively captured by the sorbent in the FCC riser. The sulfur-ladensorbent is then sent to the FCC regenerator 20 along with thecarbon-laden FCC catalyst for regeneration by treatment with anoxygen-containing gas, typically air, which is also used to regeneratethe FCC catalyst. During regeneration, sulfur carried by the sorbent isconverted to a sulfur dioxide-containing gas stream 60 that can betreated for sulfur removal in a downstream process unit such as a sulfurdioxide scrubber (not shown).

The active metal oxide sulfur sorbent has sufficient sulfur-removalactivity, and is added to the FCC reactor 110 in an amount sufficient toachieve removal of at least about 50 wt. % of sulfur contaminants whichwould otherwise be present in the FCC naphtha product, i.e., the FCCliquid product fraction having an FBP less than about 430° F.Advantageously, the active metal oxide sulfur sorbent is also active forremoval of sulfur contaminants from heavier FCC product fractions and isadded to the FCC reactor 110 in an amount sufficient to achieve removalof at least about 50 wt. % of sulfur contaminants which would otherwisebe present in both of the FCC naphtha and LCO product fractions, i.e.,the FCC liquid product fraction having an FBP of less than about 650° F.In currently preferred embodiments according to this aspect of theinvention, the active metal oxide sulfur sorbent is added to the FCCcatalyst in an amount of from about 1 to about 10 wt. %, based on theweight of the FCC catalyst.

In more preferred embodiments of this aspect of the invention, theactive metal oxide sulfur sorbent has sufficient sulfur-removalactivity, and is added to the FCC reactor 110 in an amount sufficient toachieve removal of at least about 50 wt. % of sulfur contaminants whichwould otherwise be present in the complete liquid product recovered fromthe FCC reactor. According to still other preferred embodiments, theactive metal oxide sulfur sorbent is added to the FCC reactor 110 in anamount sufficient to achieve removal of at least about 75 wt. %, morepreferably at least about 90 wt. % of sulfur contaminants which wouldotherwise be present in the naphtha product. In yet other preferredembodiments, the active metal oxide sulfur sorbent is added to the FCCreactor 110 in an amount sufficient to achieve removal of at least about75 wt. %, more preferably at least about 90 wt. % of sulfur contaminantswhich would otherwise be present in both of the FCC naphtha and LCOproduct fractions.

It has been found that regenerable sorbent comprising an active metaloxide sulfur sorbent supported on, or otherwise combined with arefractory inorganic oxide cracking catalyst are capable of removingthiophenic sulfur compounds in presence of H₂S and mercaptans. Thus,tests have shown that when a mixture of 2,000 ppmv of thiophene and10,000 ppmv of methyl mercaptan was used to test the performance of onepreferred sorbent (see Example 6), it was found that presence of 10,000ppmv of mercaptan did not affect the activity of the sorbent forthiophene removal. Similar results were also observed when thiophene wasmixed with H₂S. This is particularly important in a FCC reactor as about40 to 50% of the sulfur in the feed to the FCC is converted into H₂S. Ithas further been found that various preferred sorbents can besuccessfully regenerated under the conditions used in a typical FCCregenerator without any degradation in catalytic activity. Since thepreferred sorbents are extremely attrition-resistant, they can be usedalong with the FCC catalyst in a conventional FCC process withoutsubstantial attrition problems.

One of the added benefits of this aspect of the invention can beincreased yield of naphtha and LCO fractions from a FCC system becauseof change in sulfur distribution. Currently, refiners typically use aFBP of 410 to 420° F. for naphtha from their FCC reactor because theywant to limit the sulfur in naphtha, particularly the higher molecularweight sulfur compounds (such as alkyl dibenzothiophenes). Removal ofsulfur in the FCC riser itself, in accord with the present invention,can allow this restriction to be eased so that refiners can make premiumproducts at much higher yields than they currently do.

Although the process shown in FIG. 2 achieves desulfurization of an FCChydrocarbon feed simultaneously with the FCC process, thedesulfurization process illustrated in FIG. 2 can alternatively beachieved separately from the FCC process by treating the FCC hydrocarbonfeed in a conventional FCC unit, operated at conventional FCCconditions, and positioned upstream of the FCC processing zone.

The following examples illustrate the use of various sorbentcompositions for removal of organic sulfur compounds from varioussimulated syngas and hydrocarbon feedstocks.

EXAMPLE 1

A zinc titanate aluminate sorbent prepared according to Example 8 of PCTApplication WO 99/42201 A1, published Aug. 26, 1999, having a weight ofabout 200 g was loaded into a 2 inch ID quartz reactor. This reactor wassealed in a stainless steel pressure shell. The system was pressurizedto 50 psig and heated to 1000° F. in 4 SLPM (standard liters per minute)of nitrogen. The reactor effluent was used to continuously purge asample loop for a Varian 3300 Gas Chromatograph fitted with a SieversModel 355 sulfur chemiluminescence detector capable of detecting below200 ppbv (parts per billion, volume) of sulfur.

The test was started by adjusting the flow to the reactor to 2 SLPM ofhydrogen and 2 SLPM of a nitrogen mixture containing 200 ppmv (parts permillion volume) each of ethyl-, propyl-, and butyl-mercaptan. At thistime, HP ChemStation software was used to start a sequence designed tosample the reactor effluent at intervals of about 6 minutes. After 120minutes, the flow was adjusted to have 0.4 SLPM of hydrogen and 3.6 SLPMof the nitrogen and mercaptan mixture. At a total run time of 240minutes the flow was changed to 0.8 SLPM of 10 vol % H₂S in hydrogen and3.2 SLPM of nitrogen. When the level of H₂S in the reactor effluentreached 100 ppmv, the sulfidation was terminated.

While purging the sulfidation gases of the reactor for about 30 minuteswith 4 SLPM nitrogen, the sorbent was heated to 1150° F. After thereactor had been purged and the temperature had stabilized at the newtemperature, the sorbent was regenerated with 4 SLPM of air. Theregeneration was monitored by the SO₂ and O₂ leak in the reactoreffluent. When the O₂ level had increased above 5 vol % and the SO₂concentration had dropped below 2,000 ppmv (parts per million, volume),the regeneration was stopped.

In preparation for the next sulfidation, the sorbent bed was cooled to1000° F. Sulfidation was started with a mixture of 3.6 SLPM of hydrogen,0.2 SLPM of 1,960 ppmv thiophene in nitrogen and 0.25 SLPM of nitrogen.At the start of sulfidation, the HP ChemStation software sequenceanalyzing the reactor effluent every 6 minutes was also started. Theflows were changed to 3.6 SLPM of hydrogen, 1 SLPM of the 1,960 ppmvthiophene in nitrogen mixture and 0.25 SLPM of nitrogen after 120 min.These flow conditions were maintained for another 120 minutes. The nextset of flow conditions were 0.4 SLPM of 10 vol % H₂S in hydrogen, 3.6SLPM of hydrogen and 0.25 SLPM of nitrogen. These conditions weremaintained until the H₂S concentration in the effluent exceeded 100ppmv.

For regeneration, the sorbent bed was heated to 1150° F. Theregeneration was started with 4 SLPM of air. Regeneration was stoppedwhen the effluent SO₂ concentration dropped below 2,000 ppmv and theeffluent O₂ concentration increased above 5 vol %.

For the third sulfidation, the temperature in the sorbent bed wasdropped to 1000° F. For the first 120 minutes of sulfidation, the flowswere 3.6 SLPM of hydrogen, 0.2 SLPM of 945 ppmv 2-ethyl thiophene innitrogen and 0.3 SLPM of nitrogen. After 120 minutes, the flows werechanged to 3.6 SLPM of hydrogen, 1.0 SLPM of 945-ppmv thiophene innitrogen, and 0.3 SLPM of nitrogen. The sulfidation and, consequently,the test were then terminated. The comparison of the steady state feedand effluent concentration for the various sulfur compounds (mercaptans,thiophene and ethyl thiophene) are listed in Table 1. TABLE 1 ComparisonOf The Concentration Of The Sulfur Contaminant In The Reactor Feed AndEffluent With Zinc Titanate Aluminate Sorbent Concentration (ppmv)Compound Feed Effluent Mercaptan (Ethyl-, propyl- and butyl-) 300 0.5Mercaptan (Ethyl-, propyl- and butyl-) 540 1 Thiophene 100 1 Thiophene400 5 2-Ethylthiophene 60 0.5 2-Ethylthiophene 200 2

EXAMPLE 2

The following testing sequence was used to screen the following sorbentmaterials (1) the zinc titanate aluminate of Example 1, (2) a zincaluminate (prepared as set forth below), (3) alumina (commerciallyavailable), (4) zinc titanate, (5) a physical mixture of zinc titanateand alumina, (6) a physical mixture of zinc aluminate and zinc titanate,(7) a commercial, stabilized zinc oxide guard bed material, G72D,commercially available from Sud-Chemie Inc, and (8) ECAT, a silica basedcommercial FCC catalyst. The test began by loading 50 g of each sampleinto an 1 inch ID quartz reactor. The reactor was placed in a furnacewith temperature control based on the temperature at the center of thesorbent bed. The quartz reactor was fitted with two feed inlets, athermocouple well and effluent side arm. The reactor effluent was setupto continuously feed the sample loop of a Hewlett Packard (HP) 6890 GCfitted with a J&W GS GasPro column and a Sievers Model 355 sulfurchemiluminescence detector. This detector can easily detect sulfurconcentrations to below 200 ppbv.

In preparation for the run, the sorbent bed was heated to 800° F. in anitrogen flow of approximately 500 sccm. The test was started byintroducing into the reactor a mixture of 2,100 ppmv thiophene andnitrogen at 50 sccm (standard cubic centimeters per minute) with 400sccm of nitrogen. HP ChemStations software was used to sample thereactor effluent periodically. The reactor effluent was monitored untiltwo to three sequential results indicated steady state operation hadbeen achieved. This typically took between 40 to 60 minutes. At thispoint the reactor system was bypassed and the reactor feed was feddirectly to the GC system for analysis. As with the reactor effluent,the reactor feed was analyzed until several sequential results indicatedthe sulfur concentrations were consistent. The results from thesescreening tests are shown in Table 2.

The zinc aluminate sample used in these tests was prepared by mixing66.9 g of alumina (Engelhard) and 53.4 g of zinc oxide (Aesar) in 300 mlof deionized (DI) water. This slurry was gently heated with continuousstirring for 1 hour. The slurry was dried at 120° C. overnight andcalcined at 800° C. for 6 hours.

The effect of hydrogen addition was demonstrated in repeat test foralumina. During this test, the flows were set to 450 sccm of hydrogenand 50 sccm of a 2,100 ppmv thiophene in nitrogen mixture. The resultsfor both the test with hydrogen and without hydrogen can be seen inTable 2. TABLE 2 Comparison of Thiophene Concentration in the ReactorFeed and Effluent for Catalyst/Sorbent Screening Test Feed GasComposition Effluent N₂ H₂ Thiophene Thiophene Material (Vol %) (vol %)(ppmv) (ppmv) Zinc titanate Balance 137 114 Zinc aluminate Balance 2050.09 Alumina Balance 238 23 Alumina Balance 90.0 146 0.148 Zinc titanate(40 wt %) Balance 215 0.07 and Zinc aluminate (60 wt %) Zinc titanate(40 wt %) and Balance 195 82 alumina (60 wt %) Zinc titanate aluminateBalance 132 0.115 ECAT Balance 919 600 G72D (zinc oxide) Balance 1330.78

As can be seen in Table 2, the zinc aluminate was effective for removalof the cyclic sulfur compositions with and without added or reacted zinctitanate. Moreover, the zinc aluminate was more effective without anyhydrogen addition in removing the sulfur compounds than alumina withhydrogen. The zinc titanate aluminate was similarly effective.

EXAMPLE 3

This example used the same microreactor system that was used in Example2. An isooctane sample spiked with various sulfur compounds was used tomimic FCC naphtha (shown in Table 3). Tests were conducted with thismixture to determine the effectiveness of the zinc titanate aluminatesorbent used in Example 1 at 1,000° F. with and without H₂. The resultsare shown in Table 3. TABLE 3 Removal Of Various Sulfur Compounds From ASimulated Isooctane Sample Using Zinc Titanate Aluminate Sorbent WithAnd Without Hydrogen Product (ppmw) Feed Test 1 Test 2 Sulfur Compound(ppmw) Without H₂ With H₂ Ethyl Mercaptan 159.8 0.0 0.0 Carbon Disulfide217.7 4.7 0.0 Isopropyl Mercaptan 103.0 0.0 0.0 Thiophene 88.5 46.6 33.6Diethyl Sulfide 74.1 4.3 0.0 2-Ethyl Thiophene 62.0 54.7 43.6 DiethylDisulfide 105.1 6.6 0.8 Benzothiophene 39.8 89.8 58.3 Dibenzothiophene27.7 2.9 13.3 TOTAL 877.8 209.6 149.6 % Removal 76.1 82.9

Although not shown in Table 3, in each case the effluent was monitoredfor H₂S, and no traces were found in any of the tests. As seen in Table3, even though no hydrodesulfurization catalyst was used in any of thesetests, addition of H₂ improved the extent of desulfurization from 76.1to 82.9 percent, with significant increase in removal of benzothiopheneand dibenzothiophene. Although not fully understood, this is believeddue to the enhanced stabilization of hydrocarbon radicals resulting fromring cracking, which in turn, is believed to decrease or minimizedeactivation of the sorbent, e.g., by coking. Further, it is to be notedthat the sorbent has a surface area of about 5 m²/g, and that highersurface areas should improve the desulfurization efficiency.

EXAMPLE 4

Example 3 was repeated except that the reaction temperature was loweredto 800° F. and the zinc titanate aluminate sorbent was modified toinclude a copper promoter using the following procedure.

100 g of the zinc titanate aluminate sorbent powder of Example 3 wasdried at 120° C. for one hour and then cooled in a desiccator.

To 35 mL D.I. H₂O in a 100 ml beaker was added 28.8 g of cupric nitrate(obtained from Sigma Chemical). 5.5 mL of the Cu(NO₃)₂ solution wasapplied to the zinc titanate aluminate sorbent powder drop by drop whilestirring with a Teflon rod. The resultant powder was calcined at 200° C.(5° C./min) for 2 hours and cooled in a desiccator. The impregnation andcalcining steps were repeated to achieve a second impregnation. Thetwice impregnated sorbent was dried at 120° C. overnight, and thencalcined at 280° C. (5° C./min) for 4 hours.

The results of testing of this Cu-impregnated sorbent are shown in Table4. As can be seen from these results, the copper promoter allowed thesame sulfur removal efficiency at 800° F. as was achieved withunpromoted zinc titanate aluminate at 1000° F. TABLE 4 Removal OfVarious Sulfur Compounds With And Without The Addition Of The CopperPromoter To The Zinc Titanate Aluminate Sorbent Product (ppmw) Test 1Test 2 Feed 1,000° F. 800° F. Sulfur Compound (ppmw) (original sorbent)(modified sorbent) Ethyl Mercaptan 159.8 0.0 0.0 Carbon Disulfide 217.70.0 0.0 Isopropyl Mercaptan 103.0 0.0 0.0 Thiophene 88.5 33.6 54.6Diethyl Sulfide 74.1 0.0 175.8 2-Ethyl Thiophene 62.0 43.6 0.0 DiethylDisulfide 105.1 0.8 0.0 Benzothiophene 39.8 58.3 0.0 Dibenzothiophene27.7 13.3 0.0 TOTAL 877.8 149.6 280.4 % Removal 82.9 73.7

EXAMPLE 5

The following testing sequence was used to screen the following sorbentmaterials: (1) Iron Oxide supported on the Zinc Titanate Aluminate ofExample 1 (prepared as described below); (2) Zinc Aluminate prepared asdescribed in Example 2; (3) Copper Oxide supported on Zinc Aluminate,(prepared as described below); and, (4) Iron Oxide supported on ZincAluminate, (prepared as described below).

Preparation of sorbent (1), Iron Oxide supported on Zinc TitanateAluminate. A 100 g sample of the zinc titanate aluminate from Example 1was dried at 120° C. for an hour and allowed to cool in a desiccator. Asolution of iron nitrate was prepared by dissolving 38.3 g ofFe(NO₃)₃.9H₂O in 20 ml of deionized (DI) water. A total of 15 ml of thisiron nitrate solution was added to the zinc titanate aluminate drop bydrop while continuously mixing the zinc titanate aluminate. Theresulting powder was calcined at 200° C. for 2 hours and cooled in adesiccator. A second sample of iron nitrate solution was made andimpregnated on the previously impregnated zinc titanate aluminate in themanner described above. The final impregnated sample was dried at 120°C. overnight and calcined at 280° C. for 4 hours.

Preparation of sorbent (3), Copper Oxide supported on Zinc Aluminate. A100 g sample of the zinc aluminate from Example 2 was treated with acopper impregnating solution prepared by dissolving 44.9 g of Cu(NO₃)₂in 55 ml of DI water. During the first impregnation 26 ml of the copperimpregnating solution was added to the zinc aluminate drop by drop asthe zinc aluminate was vigorously stirred. The sample was then dried at200° C. for 2 hours and cooled in a desiccator. After cooling, thesample was impregnated with another 26 ml of the copper impregnatingsolution in the manner described above. The sample was dried at 120° C.and calcined for 4 hours at 280° C.

Preparation of sorbent (4) Iron Oxide supported on Zinc Aluminate. Aniron impregnated zinc aluminate sample was prepared using the sameprocedure as used for the copper impregnated zinc aluminate of sorbent(3) above. The iron impregnating solution was prepared by dissolving76.2 g of Fe(NO₃)₃.9H₂O in 40 ml of DI water. The twice impregnatedsample was dried and calcined in a like manner as sorbent (3) above.

The test began by loading 50 g of each sample into a 1-inch ID quartzreactor. The reactor was placed in a furnace with temperature controlbased on the temperature at the center of the sorbent bed. The quartzreactor was fitted with two feed inlets, a thermocouple well, and aneffluent side arm. The reactor effluent was setup to continuously feedthe sample loop of a HP 6890 GC fitted with a J&W GC GasPro column and aSievers Model 355 sulfur chemiluminescence detector. This detector caneasily detect sulfur down to 50 ppbv.

In preparation for each test, the sorbent bed was heated to 800° F. in anitrogen flow of approximately 500 sccm. The test was started byintroducing into the reactor a mixture containing 200 ppmvmethylmercaptan, and 200 ppmv thiophene with the balance being nitrogen.HP Chemstations software was used to sample the reactor effluentperiodically. The reactor effluent was monitored until two or threesequential results indicated steady state operation had been achieved.This typically took between 40 to 60 minutes. At this point the reactorsystem was bypassed and the reactor feed was feed directly to the GCsystem for analysis. As with the reactor effluent, the reactor feed wasanalyzed until several sequential results indicated the sulfurconcentrations were consistent. The results from these screening testsare shown in Table 5. TABLE 5 Comparison of Reactor Feed and EffluentFor Second Sorbent Screening Test Methyl Mercaptan Thiophene (ppmv)(ppmv) Sorbent Material Feed Effluent Feed Effluent Iron Oxide/ZincTitanate Aluminate 186  N.D.* 274 N.D. Zinc aluminate 191 N.D. 281 0.7Copper Oxide/Zinc Aluminate 191 N.D. 290 N.D. Iron Oxide/Zinc Aluminate191 N.D. 291 0.2*Not Detected

EXAMPLE 6

A 50 g sample of the Zinc Aluminate-supported Iron Oxide sorbentprepared as described in Example 5 was loaded in the 1-inch ID quartzreactor. The furnace heating was controlled with a thermocouple in thesorbent bed approximately 1-in from the quartz frit supporting thesorbent bed. After installing the quartz reactor and connecting the feedand effluent lines, the sorbent bed was heated to 800° F. in a nitrogenflow of approximately 500 sccm. When the sorbent bed temperature was800° F., the sorbent was exposed to 500 sccm of air for 60 min. Thereactor was purged with nitrogen at 500 sccm for 15 min to remove anytraces of oxygen. The sample was then exposed to a mixture with 1920ppmv of thiophene and 9940 ppmv methyl mercaptan in nitrogen at 500sccm. HP Chemstations software was used to periodically record thesulfur content of the reactor effluent as determined by an HP 6890 GCequipped with a J&W GasPro column and Sievers Model 355 sulfurchemiluminescence detector. Exposure of the sorbent sample continueduntil the thiophene concentration in the effluent increased to 100 ppmv.At this point no methyl mercaptan was detected in the effluent. Thetotal time of sorbent exposure prior to breakthrough (thiophene effluentconcentration >100 ppmv) was 5 hours. This corresponds to a sulfurweight loading of 4.4 wt % for the methtyl mercaptan and 0.7 wt % forthe thiophene.

The sorbent sample was then regenerated with 500 sccm of air at 800° F.for 60 min. The sorbent was exposed to the same methyl mercaptan,thiophene and nitrogen mixture at the same conditions as during thefirst exposure to breakthrough. The total exposure time prior tobreakthrough for this second exposure was 4 hours. Once again thethiophene effluent concentration was observed to increase to 100 ppmvwithout any methyl mercaptan being detected. The sulfur loadings were0.84 wt % for thiophene and 3.6 wt % for methyl mercaptan.

The sorbent was again regenerated with 500 sccm of air at 800° F. for120 min. After purging of the oxygen by nitrogen, the sorbent wasexposed to a 1970 ppmv thiophene in nitrogen mixture at 500 sccm at 800°F. The effluent sulfur content was monitored as in previous exposurecycles. The sorbent was exposed to this mixture for 6 hours. The testhad to be terminated at this point because the tank with thethiophene/nitrogen mixture was empty. The effluent thiopheneconcentration at this time was 56 ppmv. Thus, breakthrough had not beenreached. The sulfur loading for this exposure test was 1 wt % forthiophene.

Many modifications and other embodiments of the invention will come tomind to one skilled in the art to which this invention pertains havingthe benefit of the teachings presented in the foregoing descriptions andthe associated drawing. Therefore, it is to be understood that theinvention is not to be limited to the specific embodiments disclosed andthat modifications and other embodiments are intended to be includedwithin the spirit and scope of the appended claims. Although specificterms are employed herein, they are used in a generic and descriptivesense only and not for purposes of limitation.

1. A process for removing sulfur compounds from a normally liquidhydrocarbon fuel or fuel component feedstock having a sulfur content ofat least about 150 ppmw comprising the steps: contacting the feedstockin the substantial absence of a hydrodesulfurization catalyst, with aregenerable sorbent material comprising at least one active metal oxidesorbent capable of selectively removing sulfur compounds present in thehydrocarbon feedstock and a refractory inorganic oxide cracking catalystcapable of cracking cyclic organic sulfur compounds; and recovering ahydrocarbon product having a sulfur content of about 50% or less thanthe sulfur content of the feedstock.
 2. The process of claim 1, furthercomprising regenerating at least a portion of said sorbent with anoxidizing gas under conditions sufficient to convert metal sulfide intosaid metal oxide sorbent and thereby provide regenerated sorbent, andrecycling at least a portion of said regenerated sorbent to saidcontacting step.
 3. The process of claim 1, wherein said refractoryinorganic oxide cracking catalyst comprises at least onemetal-substituted refractory inorganic oxide cracking catalyst, saidmetal being the same metal as the metal of said active metal oxidesorbent.
 4. The process of claim 1, wherein said contacting step isconducted at a temperature of at least about 300° C.
 5. The process ofclaim 1, wherein said hydrocarbon feedstock comprises at least about 100ppmw of cyclic organic sulfur compounds.
 6. The process of claim 5,wherein said wherein said hydrocarbon feedstock comprises a sulfurcontent of at least about 300 ppmw.
 7. The process of claim 1 whereinsaid contacting step is conducted such that said feedstock is contactedsimultaneously with said sorbent and said refractory inorganic oxidecracking catalyst.
 8. The process of claim 2, further comprisingregenerating at least a portion of said refractory inorganic oxidecracking catalyst with an oxidizing gas under conditions sufficient toremove sulfur from said refractory inorganic oxide cracking catalyst andthereby provide regenerated refractory inorganic oxide crackingcatalyst, and recycling at least a portion of said regeneratedrefractory inorganic oxide cracking catalyst to said contacting step. 9.The process of claim 1, wherein said hydrocarbon feedstock comprises FCCnaphtha.
 10. The process of claim 1, wherein said hydrocarbon feedstockconsists essentially of FCC naphtha.
 11. The process of claim 9, whereinsaid hydrocarbon product recovered in said recovering step has a sulfurcontent of less than about 10 ppmw.
 12. The process of claim 1, whereinsaid hydrocarbon feedstock comprises diesel fuel or a precursor orcomponent thereof.
 13. The process of claim 12, wherein said hydrocarbonfeedstock comprises coker naphtha, thermally cracked naphtha, lightcycle oil, or a straight-run diesel fraction.
 14. The process of claim1, wherein said metal oxide sorbent comprises zinc oxide.
 15. Theprocess of claim 1, wherein said refractory inorganic oxide crackingcatalyst comprises alumina or a metal-substituted alumina.
 16. Theprocess of claim 1, wherein said metal oxide sorbent comprises metaloxide sorbent and said refractory inorganic oxide cracking catalystcomprise zinc oxide and zinc aluminate.
 17. The process of claim 1,wherein said contacting step is carried out in a transport bed reactorwith a vapor residence time of less than about 20 seconds.
 18. A processfor removing cyclic and polycyclic organic sulfur compounds from anormally liquid hydrocarbon feedstock comprising the steps: contactingthe feedstock in the substantial absence of a hydrodesulfurizationcatalyst, with a sorbent comprising a metal-substituted refractoryinorganic oxide cracking catalyst capable of cracking cyclic organicsulfur compounds, said metal being selected from the group consisting ofmetals which are capable in their oxide form, of adsorption of reducedsulfur compounds by conversion of the metal oxide to a metal sulfide;and recovering a hydrocarbon product having a cyclic and polycyclicorganic sulfur content at least about 25% less than the cyclic andpolycyclic organic sulfur content of the feedstock, based the sulfurweight of said cyclic and polycyclic organic sulfur compounds in saidfeedstock and the sulfur weight of cyclic and polycyclic organic sulfurcompounds in said product.
 19. The process of claim 18, furthercomprising regenerating at least a portion of said sorbent with anoxidizing gas under conditions sufficient to convert metal sulfide intosaid metal oxide and thereby provide regenerated sorbent, and recyclingat least a portion of said regenerated sorbent to said contacting step.20. The process of claim 18, wherein said sorbent further comprises anactive metal oxide sorbent capable of selectively removing sulfurcompounds present in the hydrocarbon feedstock, the metal of said metaloxide being the same metal as the metal of said metal-substitutedrefractory inorganic oxide cracking catalyst sorbent.
 21. The process ofclaim 18, wherein said contacting step is conducted at a temperature ofat least about 300° C.
 22. The process of claim 18, wherein saidhydrocarbon feedstock comprises at least about 150 ppmw of sulfurcompounds.
 23. The process of claim 18, wherein said product has asulfur content at least about 50% less than the sulfur content of thefeedstock.
 24. The process of claim 23, wherein said hydrocarbonfeedstock comprises FCC naphtha.
 25. The process of claim 18, whereinsaid hydrocarbon feedstock comprises FCC naphtha.
 26. The process ofclaim 23, wherein said hydrocarbon feedstock consists essentially of FCCnaphtha.
 27. The process of claim 18, wherein said hydrocarbon feedstockconsists essentially of FCC naphtha.
 28. The process of claim 24,wherein said hydrocarbon product has a sulfur content of less than about10 ppmw.
 29. The process of claim 18, wherein said hydrocarbon feedstockcomprises diesel fuel or a precursor or component thereof.
 30. Theprocess of claim 18, wherein said hydrocarbon feedstock consistsessentially of diesel fuel or a precursor or component thereof.
 31. Theprocess of claim 29, wherein said hydrocarbon feedstock comprises cokernaphtha, thermally cracked naphtha, light cycle oil, or a straight-rundiesel fraction.
 32. The process of claim 30, wherein said hydrocarbonfeedstock comprises coker naphtha, thermally cracked naphtha, lightcycle oil, or a straight-run diesel fraction.
 33. The process of claim18, wherein said metal-substituted refractory inorganic oxide crackingcatalyst comprises zinc aluminate.
 34. The process of claim 20, whereinsaid metal-substituted refractory inorganic oxide cracking catalystcomprises zinc aluminate.
 35. The process of claim 20, wherein saidactive metal oxide sorbent comprises zinc oxide.
 36. The process ofclaim 28, wherein said active metal oxide sorbent comprises zinctitanate.
 37. The process of claim 18, wherein said metal-substitutedrefractory inorganic oxide cracking catalyst comprises iron aluminate.38. The process of claim 20, wherein said metal-substituted refractoryinorganic oxide cracking catalyst comprises iron aluminate.
 39. Theprocess of claim 20, wherein said active metal oxide sorbent comprisesan iron oxide.
 40. The process of claim 18, wherein said contacting stepis carried out in a transport bed reactor with a vapor residence time ofless than about 20 seconds.
 41. The process of claim 18, wherein saidcontacting step is carried out in a bubbling bed reactor.
 42. Theprocess of claim 20, wherein said contacting step is carried out in atransport bed reactor with a vapor residence time of less than about 20seconds.
 43. The process of claim 20, wherein said contacting step iscarried out in a bubbling bed reactor.
 44. The process of claim 24,wherein said contacting step is carried out in a transport bed reactorwith a vapor residence time of less than about 20 seconds.
 45. Theprocess of claim 24, wherein said contacting step is carried out in abubbling bed reactor.
 46. The process of claim 29, wherein saidcontacting step is carried out in a transport bed reactor with a vaporresidence time of less than about 20 seconds.
 47. The process of claim29, wherein said contacting step is carried out in a bubbling bedreactor.
 48. A process for removing cyclic and polycyclic organic sulfurcompounds from a normally liquid hydrocarbon feedstock comprising thesteps: contacting the feedstock in the substantial absence of ahydrodesulfurization catalyst, with a sorbent comprising at least oneactive metal oxide sorbent capable of selectively removing sulfurcompounds present in the hydrocarbon feedstock and a refractoryinorganic oxide cracking catalyst capable of cracking cyclic organicsulfur compounds; and recovering a hydrocarbon product having a cyclicand polycyclic organic sulfur content at least about 25% less than thecyclic and polycyclic organic sulfur content of the feedstock, based thesulfur weight of said cyclic and polycyclic organic sulfur compounds insaid feedstock and the sulfur weight of cyclic and polycyclic organicsulfur compounds in said product.
 49. The process of claim 48, furthercomprising regenerating at least a portion of said sorbent with anoxidizing gas under conditions sufficient to convert metal sulfide intosaid metal oxide and thereby provide regenerated sorbent, and recyclingat least a portion of said regenerated sorbent to said contacting step.50. The process of claim 48, wherein said contacting step is conductedat a temperature of at least about 300° C.
 51. The process of claim 48,wherein said hydrocarbon feedstock comprises at least about 150 ppmw ofsulfur compounds.
 52. The process of claim 48, wherein said product hasa sulfur content at least about 50% less than the sulfur content of thefeedstock.
 53. The process of claim 52, wherein said hydrocarbonfeedstock comprises FCC naphtha.
 54. The process of claim 48, whereinsaid hydrocarbon feedstock comprises hydrotreated FCC naphtha.
 55. Theprocess of claim 48, wherein said hydrocarbon feedstock compriseshydrotreated diesel fuel or a hydrotreated precursor or hydrotreatedcomponent thereof.
 56. The process of claim 48, wherein said hydrocarbonfeedstock consists essentially of a hydrotreated gasoline or diesel fuelor a hydrotreated precursor or hydrotreated component of gasoline ordiesel fuel.
 57. The process of claim 56, wherein said hydrocarbonproduct has a sulfur content of less than about 10 ppmw.
 58. The processof claim 48, wherein said hydrocarbon feedstock comprises diesel fuel ora precursor or component thereof.
 59. The process of claim 48, whereinsaid hydrocarbon feedstock consists essentially of diesel fuel or aprecursor or component thereof
 60. The process of claim 59, wherein saidhydrocarbon feedstock comprises coker naphtha, thermally crackednaphtha, light cycle oil, or a straight-run diesel fraction.
 61. Theprocess of claim 52, wherein said hydrocarbon feedstock comprises cokernaphtha, thermally cracked naphtha, light cycle oil, or a straight-rundiesel fraction.
 62. The process of claim 48, wherein said refractoryinorganic oxide cracking catalyst comprises zinc aluminate.
 63. Theprocess of claim 48, wherein said refractory inorganic oxide crackingcatalyst comprises iron aluminate.
 64. The process of claim 48, whereinsaid active metal oxide sorbent comprises zinc oxide.
 65. The process ofclaim 48, wherein said active metal oxide sorbent comprises an ironoxide.
 66. The process of claim 48, wherein said contacting step iscarried out in a transport bed reactor with a vapor residence time ofless than about 20 seconds.
 67. The process of claim 48, wherein saidcontacting step is carried out in a bubbling bed reactor.
 68. Theprocess of claim 52, wherein said contacting step is carried out in atransport bed reactor with a vapor residence time of less than about 20seconds.
 69. The process of claim 52, wherein said contacting step iscarried out in a bubbling bed reactor.
 70. A process for removing sulfurcompounds from a normally liquid hydrocarbon fuel or fuel componentfeedstock having a sulfur content of at least about 150 ppmw comprisingthe steps: contacting the feedstock in a transport bed reactor during avapor residence time of less than about 20 seconds, with a regenerablesorbent material comprising at least one active metal oxide sorbentcapable of selectively removing sulfur compounds present in thehydrocarbon feedstock and a refractory inorganic oxide cracking catalystcapable of cracking cyclic organic sulfur compounds, said reactor beingsubstantially free of hydrodesulfurization catalyst; and recovering ahydrocarbon product having a reduced sulfur content.
 71. The process ofclaim 70, further comprising regenerating at least a portion of saidsorbent with an oxidizing gas under conditions sufficient to convertmetal sulfide into said metal oxide sorbent and thereby provideregenerated sorbent, and recycling at least a portion of saidregenerated sorbent to said contacting step.
 72. The process of claim70, wherein said refractory inorganic oxide cracking catalyst comprisesat least one metal-substituted refractory inorganic oxide crackingcatalyst, said metal being the same metal as the metal of said activemetal oxide sorbent.
 73. The process of claim 70, wherein saidcontacting step is conducted at a temperature of at least about 300° C.74. The process of claim 70, wherein said hydrocarbon feedstockcomprises at least about 100 ppmw of cyclic and polycyclic organicsulfur compounds.
 75. The process of claim 70, wherein said wherein saidhydrocarbon feedstock comprises a sulfur content of at least about 300ppmw.
 76. The process of claim 70 wherein said contacting step isconducted such that said feedstock is contacted simultaneously with saidsorbent and said refractory inorganic oxide cracking catalyst.
 77. Theprocess of claim 72, further comprising regenerating at least a portionof said refractory inorganic oxide cracking catalyst with an oxidizinggas under conditions sufficient to remove sulfur from said refractoryinorganic oxide cracking catalyst and thereby provide regeneratedrefractory inorganic oxide cracking catalyst, and recycling at least aportion of said regenerated refractory inorganic oxide cracking catalystto said contacting step.
 78. The process of claim 70, wherein saidhydrocarbon feedstock comprises FCC naphtha.
 79. The process of claim70, wherein said hydrocarbon feedstock comprises diesel fuel or aprecursor or component thereof.
 80. The process of claim 70, whereinsaid hydrocarbon product recovered in said recovering step has a sulfurcontent of less than about 10 ppmw.
 81. The process of claim 70, whereinsaid metal oxide sorbent comprises zinc oxide.
 82. The process of claim70, wherein said refractory inorganic oxide cracking catalyst comprisesalumina or a metal-substituted alumina.
 83. The process of claim 70,wherein said metal oxide sorbent comprises an iron oxide.
 84. Theprocess of claim 70, wherein said refractory inorganic oxide crackingcatalyst comprises iron aluminate.
 85. A process for removing cyclic andpolycyclic organic sulfur compounds from a normally liquid hydrocarbonfeedstock having a sulfur content comprising at least about 100 ppmw ofcyclic and polycyclic organic sulfur compounds comprising the steps:contacting the feedstock in a transport bed reactor during a vaporresidence time of less than about 20 seconds with a sorbent comprising ametal-substituted refractory inorganic oxide cracking catalyst capableof cracking cyclic organic sulfur compounds, said metal being selectedfrom the group consisting of metals which are capable in their oxideform, of adsorption of reduced sulfur compounds by conversion of themetal oxide to a metal sulfide, said reactor being substantially free ofhydrodesulfurization catalyst; and recovering a hydrocarbon producthaving a cyclic and polycyclic organic sulfur content at least about 25%less than the cyclic and polycyclic organic sulfur content of thefeedstock, based the sulfur weight of said cyclic and polycyclic organicsulfur compounds in said feedstock and the sulfur weight of cyclic andpolycyclic organic sulfur compounds in said product.
 86. The process ofclaim 85, further comprising regenerating at least a portion of saidsorbent with an oxidizing gas under conditions sufficient to convertmetal sulfide into said metal oxide and thereby provide regeneratedsorbent, and recycling at least a portion of said regenerated sorbent tosaid contacting step.
 87. The process of claim 85, wherein said sorbentfurther comprises an active metal oxide sorbent capable of selectivelyremoving sulfur compounds present in the hydrocarbon feedstock, saidmetal being the same metal as the metal of said metal-substitutedrefractory inorganic oxide cracking catalyst sorbent.
 88. The process ofclaim 85, wherein said contacting step is conducted at a temperature ofat least about 300° C.
 89. The process of claim 85, wherein saidhydrocarbon feedstock comprises at least about 300 ppmw of sulfurcompounds.
 90. The process of claim 86, wherein said wherein saidproduct has a sulfur content at least about 50% less than the sulfurcontent of the feedstock.
 91. The process of claim 85, wherein saidhydrocarbon feedstock comprises an FCC naphtha.
 92. The process of claim85, wherein said hydrocarbon feedstock comprises diesel fuel or aprecursor or component thereof.
 93. The process of claim 85, whereinsaid hydrocarbon product recovered in said recovering step has a sulfurcontent of less than about 10 ppmw.
 94. The process of claim 87, whereinsaid metal oxide sorbent comprises zinc oxide.
 95. The process of claim85, wherein said metal-substituted refractory inorganic oxide crackingcatalyst comprises a metal-substituted alumina.
 96. The process of claim87, wherein said metal oxide sorbent comprises an iron oxide.
 97. Theprocess of claim 85, wherein said metal-substituted refractory inorganicoxide cracking catalyst comprises iron aluminate.
 98. A process forremoving organic sulfur compounds from an FCC hydrocarbon stream duringan FCC process comprising the steps: contacting an FCC hydrocarbonfeedstock in a reaction zone under FCC reaction conditions with an FCCcatalyst and a regenerable sorbent comprising an active metal oxidesulfur sorbent supported on or otherwise combined with a refractoryinorganic oxide cracking catalyst, said metal being selected from thegroup consisting of metals which are capable in their oxide form, ofadsorption of reduced sulfur compounds by conversion of the metal oxideto a metal sulfide; and recovering a cracked hydrocarbon productcomprising FCC naphtha having a sulfur content at least about 50 wt. %less than the sulfur content of said FCC naphtha when said FCC processis conducted without said regenerable sorbent under substantiallyidentical FCC reaction conditions.
 99. The process of claim 98, furthercomprising regenerating at least a portion of said sorbent and said FCCcatalyst with an oxidizing gas under FCC catalyst regeneratingconditions to thereby remove sulfur from said sorbent and therebyregenerate said sorbent and said FCC catalyst, and recycling at least aportion of the regenerated sorbent and regenerated FCC catalyst said tosaid contacting step.
 100. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step comprises FCCnaphtha and light cycle oil fractions having a sulfur content at leastabout 50 wt. % less than the sulfur content of said FCC naphtha andlight cycle oil fractions when said FCC process is conducted withoutsaid regenerable sorbent under substantially identical FCC reactionconditions.
 101. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step comprises FCCnaphtha having a sulfur content at least about 75 wt. % less than thesulfur content of said FCC naphtha when said FCC process is conductedwithout said regenerable sorbent under substantially identical FCCreaction conditions.
 102. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step comprises FCCnaphtha having a sulfur content at least about 90 wt. % less than thesulfur content of said FCC naphtha when said FCC process is conductedwithout said regenerable sorbent under substantially identical FCCreaction conditions
 103. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step has a sulfurcontent at least about 50 wt. % less than the sulfur content of saidcracked hydrocarbon product when said FCC process is conducted withoutsaid regenerable sorbent under substantially identical FCC reactionconditions.
 104. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step comprises FCCnaphtha and light cycle oil fractions having a sulfur content at leastabout 75 wt. % less than the sulfur content of said FCC naphtha andlight cycle oil fractions when said FCC process is conducted withoutsaid regenerable sorbent under substantially identical FCC reactionconditions.
 105. The process of claim 98 wherein said crackedhydrocarbon product recovered in said recovering step comprises FCCnaphtha and light cycle oil fractions having a sulfur content at leastabout 90 wt. % less than the sulfur content of said FCC naphtha andlight cycle oil fractions when said FCC process is conducted withoutsaid regenerable sorbent under substantially identical FCC reactionconditions.
 106. The process of claim 98 wherein regenerable sorbent ispresent in said reaction zone an amount of from about 1 to about 10 wt%, based on the weight of the FCC catalyst present in said reactionzone.
 107. The process of claim 98, wherein said a refractory inorganicoxide cracking catalyst consists essentially of a metal-substitutedrefractory inorganic oxide cracking catalyst.
 108. The process of claim107, wherein said the metal of said active metal oxide sulfur sorbent isthe same metal as the metal of said metal-substituted refractoryinorganic oxide cracking catalyst sorbent.
 109. The process of claim108, wherein said metal-substituted refractory inorganic oxide crackingcatalyst comprises zinc aluminate.
 110. The process of claim 107,wherein said metal-substituted refractory inorganic oxide crackingcatalyst comprises zinc aluminate.
 111. The process of claim 98, whereinsaid active metal oxide sulfur sorbent comprises zinc oxide.
 112. Theprocess of claim 98, wherein said active metal oxide sulfur sorbentcomprises zinc titanate.
 113. The process of claim 107, wherein saidmetal-substituted refractory inorganic oxide cracking catalyst comprisesiron aluminate.
 114. The process of claim 108, wherein saidmetal-substituted refractory inorganic oxide cracking catalyst comprisesiron aluminate.
 115. The process of claim 98, wherein said active metaloxide sulfur sorbent comprises an iron oxide.